The utility of numerous refractoring oxides, such as silica, alumina, zironia, magnesia, beryllia, amorphous and crystalline aluminosilicates and the like for hydrocarbon conversion is well known. However, as is the case in every process involving expensive raw materials, catalysts and operating equipment, there remains considerable room for improvement in every aspect of these procedures. For an example, many hydrocarbon feeds contain organonitrogen compounds -- known catalyst inhibitors -- in such quantitites that it is generally preferable, and often essential, to reduce the organonitrogen content before attempting to effect other conversions such as cracking, hydrocracking, isomerization, or related processes.
Every catalyst presently employed for this purpose requires the use of relatively expensive materials in its composition, such as refractory oxide supports and hydrogenation components. They further require that the feed be contacted in costly pressurized equipment at elevated temperatures for periods sufficient to effect the desired reduction in organonitrogen content. The development of systems which can do this with less expensive compositions or lesser amounts thereof, less severe reaction conditions and shorter contact times is obviously a worthy objective.
Another area of hydrocarbon conversion in which there also remains room for improvement is that of hydrocracking, particularly midbarrel hydrocracking. In the hydrogenative conversion of hydrocarbons to midbarrel range products boiling between about 300.degree. and 700.degree. F. there is always some conversion of feed constituents to products boiling below the desired minimum product boiling point. In fact a substantial proportion of the feed is always converted to very low molecular weight products referred to as "dry gas." Production of these and other low molecular weight materials reduces midbarrel yield and is thus undesirable. Even a minor improvement in selectivity can result in substantial savings by minimizing by-products formation and the expense of separating and handling those by-products.
For example, if a hydrocracking process intended for the production of midbarrel fuels is operated at 50% conversion per pass and 50% selectivity to the desired boiling range product, an increase in selectivity of only 5% can increase yields of the desired product by 10% on a relative basis. On a commercial scale wherein the average unit usually consumes approximately 20,000 barrels a day of feedstock, this difference in selectivity can result in a savings of roughly 2,000 barrels of feed a day.
I have now discovered a procedure involving a particular catalyst composition whereby hydrocarbons can be converted more efficiently to products boiling within a predetermined boiling range. I have also discovered a procedure whereby the denitrogenation activity of certain catalyst compositions can be markedly improved.
It is therefore one object of this invention to provide an improved hydrocarbon conversion catalyst. It is another object to provide a catalyst exhibiting higher selectivity to products boiling within a predetermined boiling range. Another object is the provision of a method for producing such catalyst. Yet another object involves the provision of a midbarrel hydrocracking process having higher selectivity to midbarrel fuels. Another object is the provision of an improved hydrocarbon denitrogenation catalyst. Another objective involves the provision of a more efficient hydrocarbon denitrogenation process.
In accordance with one embodiment of this invention an improved hydrocarbon conversion catalyst containing at least one amorphous refractory oxide, a catalytically active amount of one or more crystalline zeolitic aluminosilicates and a hydrogenation component selected from the Group VI and VIII metals, oxides and sulfides is prepared by thermally activating the amorphous oxide at a temperature of at least about 600.degree. F., preferably about 600.degree. to about 1400.degree. F., intimately admixing it with the aluminosilicate in finely divided form, forming a particle aggregate of the oxide, zeolite and hydrogenation component or hydrogenation component precursor, and thermally activating the resultant aggregate. Such activation is preferably effected at a temperature of at least about 600.degree. F.
The combination of the oxide and crystalline aluminosilicate can be dried if desired, e.g., at a temperature below 600.degree. F., prior to combination with the hydrogenation component. However, such treatment is not essential. In fact, it is more economical to defer any thermal treatment whatever until addition of the hydrogenation components. Accordingly, the preactivated oxide, zeolite and hydrogenation components are preferably formed into an aggregate mass, such as a pellet, tablet, extrudate, coating or the like without intermediate thermal treatment.
This procedure leads to the formation of a catalyst having several improved characteristics, notably improved denitrogenation activity and hydrocracking selectivity to a predetermined boiling range product, particularly midbarrel range products. As a result of this observation, I have found that more desirable catalysts and methods employing the same can be obtained without the expense or complexity of more involved alternative procedures. For example, the final catalyst could be produced by first forming an aggregate of the aluminosilicate and alumina, thermally activating that aggregate and then adding the hydrogenation component. The procedures and compositions herein disclosed eliminate the need for this multi-step process and enable the production of highly active and selective compositions by much simpler means.
The presently preferred method involves intimately admixing the oxide, zeolite and hydrogenation component with sufficient water to form a paste suitable for extrusion, pelleting or the like. The amount of aqueous medium added during the admixture of these materials is preferably sufficient only to produce a formable paste. This procedure eliminates the need for cumbersome separation, drying or other steps necessary to remove relatively large quantities of excess water.
In accordance with another embodiment of this invention, I have discovered that organonitrogen containing feedstocks can be more efficiently denitrogenated by contacting with hydrogen with the above-described catalyst under denitrogenation conditions of temperature, pressure and contact time as hereinafter detailed. As a result of this method the rate of organonitrogen conversion can be increased several fold so that the same degree of denitrogenation can be obtained under much less severe conditions or with shorter contact times. Conversely higher denitrogenation rates are realized at otherwise identical conditions.
Another embodiment involves an improved hydrocracking method whereby markedly higher relative conversions to products boiling with a predetermined boiling range are obtained. In particular I have discovered that the specific combination of reaction conditions, catalysts and feed compositions hereinafter detailed affords higher hydrocracking selectivity to lower molecular weight products boiling within a prescribed range. For instance under otherwise identical conditions these systems are capable of producing 10 relative percent more midbarrel range products, boiling between 300.degree. and 700.degree. F. than are alternative methods.
Essentially any crystalline zeolitic aluminosilicate can be employed in these compositions. A preferred class of aluminosilicates includes the crystalline species having SiO.sub.2 /Al.sub.2 O.sub.3 ratios of at least about 2. This class includes both synthetic and naturally occurring zeolites. Illustrative of the synthetic zeolites are Zeolite X, U.S. Pat. Nos. 2,882,244; Zeolite Y, 3,130,007; Zeolite A, 2,882,243; Zeolite L, Bel. 575,117; Zeolite D, Can. 611,981; Zeolite R, 3,030,181; Zeolite S, 3,054,657; Zeolite T, 2,950,952; Zeolite Z, Can. 614,995; Zeolite E, Can. 636,931; Zeolite F, 2,995,358; Zeolite O, 3,140,252; Zeolite B, 3,008,803; Zeolite Q, 2,991,151; Zeolite M, 2,995,423; Zeolite H, 3,010,789; Zeolite J, 3,001,869; Zeolite W, 3,012,853; Zeolite KG, 3,056,654; Zeolite SL, Dutch 6,710,729; Zeolite Omega, Can. 817,915; synthetic mordenite; the so-called ultrastable zeolites of U.S. Pat. Nos. 3,293,192 and 3,449,070; the so-called layered aluminosilicates such as those described in U.S. Pat. Nos. 3,252,757 and 3,252,889, and the like. Illustrative of the naturally occurring crystalline zeolites are levynite, dachiardite, erionite, faujasite, analcite, paulingite, noselite, ferrierite, haulandite, scolecite, stilbite, clinoptilolite, harmotone, phillipsite, brewsterite, flakite, datolite, chabazite, gmelinite, cancrinite, leucite, lazurite, scolecite, mesolite, ptilolite, mordenite, nepheline, natrolite. Zeolites which are presently most preferred include the synthetic faujasites X and Y, zeolite T, L omega, mordenite and pretreated and post treated forms thereof such as the acid extracted and so-called ultrastable zeolites. This preference is due primarily to chemical and physical properties such as pore size, pore volume, surface area, ion exchange capacity, physical and chemical stability and catalytic activity.
Although the advantages of this invention can be realized with the foregoing aluminosilicates, I presently prefer the use of pretreated zeolites having exceptionally high thermal, hydrothermal and reammoniation stability, activity and selectivity as hereinafter described. In accordance with this preferred embodiment the starting material is first exchanged with hydrogen ions or hydrogen ions precursors in amounts sufficient to occupy at least 20 percent of the ion exchange capacity of the zeolite. A corresponding amount of the alkali metal originally present in the zeolite is replaced by the hydrogen ion precursors or hydrogen ions introduced by direct exchange. This first exchange step is preferably sufficient to reduce the alkali metal content to less than 3 percent, preferably less than 2 percent. This procedure is usually sufficient to introduce at least about 0.5 milliequivalents of hydrogen ions or hydrogen ion precursors per gram of zeolite. Of course, each of these exchanges can be carried out in a single step or a plurality of steps, the latter approach often being preferred or even necessary to obtain the desired degree of exchange.
Hydrogen ion precursors are generally well known and include ions which are exchangeable into aluminosilicates and decompose upon exposure to elevated temperatures to form the hydrogen or decationized zeolite. Illustrative of these materials are the organic and inorganic ammonium salts such as ammonium halides, e.g., chlorides, bromides, ammonium carbonates, ammonium thicynate, ammonium hydroxide, ammonium molybdate, ammonium dithionate, ammonium nitrate, ammonium sulfate, ammonium formate, ammonium lactate, ammonium tartrate and the like. Other suitable exchange compounds include the class of organic nitrogen bases such as pyridine, guanidine, and quinoline salts. Another class of organic compounds includes the complex polyhydrocarbyl ammonium salts, e.g., the tri- and tetraalkyl and aryl salts such as trimethylammonium hydroxide and tetraethylammonium hydroxide.
In the alternative the hydrogen ion can be introduced directly in the first exchange step by contacting the aluminosilicate with a hydrogen ion donor such as an organic or inorganic acid. Hydrogen ions introduced in this manner are herein referred to as unstabilized hydrogen ions since they have not yet been subjected to stabilizing thermal treatment. Illustrative inorganic acids include hydrochloric, phosphoric, sulfuric, nitric, sulfurous, chloroplatinic, dithionic, thiocyanic, carbonic, nitrous and the like. The organic acids include the mono-, di- and poly-carboxylic acids having either aliphatic, cycloaliphatic or aromatic hydrocarbyl radicals. Illustrative of these compounds are formic acid, propionic acid, melanic acid, alkenylsuccinic acid, itaconic acid, malonic acid, acetic acid, chloroacetic acid, 1,4-cyclohexadicarboxylic acid, terephthalic acid, 1,8-naphthalenedicarboxylic acid, 3-carboxycinnamic acid, phenylacetic acid, benzoic acid, substituted aromatic acids such as the chlorohydroxy-, or nitro-substituted benzoic acid and the like. However, it is presently preferred that the hydrogen ion be introduced by exchange with an inorganic ammonium salt such as ammonium nitrate or ammonium sulfate and thermal conversion to hydrogen ion.
I have found that in order to produce a composition having the desired ultimate properties it is essential that the zeolite be steamed following the first exchange, as opposed to calcination under anhydrous conditions. It is believed that maintaining at least a measurable amount of water vapor in the vicinity of the zeolite during this first thermal treatment is necessary to preserve a higher degree of the structural integrity while maintaining ion exchange capacity, catalytic activity, increasing pore size distribution and improving selectivity to midbarrel fuels under hydrocracking conditions. Accordingly, this thermal treatment is usually conducted in the presence of at least about 0.2, usually at least 2 and preferably about 5 to about 15 psi water vapor partial pressure.
The zeolite can be steamed by any procedure capable of maintaining a substantial amount of water vapor in the presence of the zeolite during at least the initial stages of the thermal treatment. For example, the exchanged zeolite can be introduced into a batch or continuous rotary furnace, a moving bed furnace or static bed calcination zone into which humidified air, or more preferably pure steam, is introduced either concurrently or countercurrently. In the alternative, water vapor released by the zeolite during the initial stages of calcination can be trapped and retained in the presence of the zeolite.
Steaming should be effected at a temperature sufficient to thermally stabilize and/or convert the zeolite to the corresponding hydrogen or decationized form yet insufficient to thermally degrade a substantial portion of the aluminosilicate structure. Steaming temperatures are usually in excess of 600.degree. F., preferably about 800.degree. to about 1650.degree. F. The zeolite is subjected to these temperatures for a period sufficient to convert it to the stabilized hydrogen form. The duration of this treatment is usually at least about 0.5 minutes, preferably about 30 minutes to about 4 hours at temperature. Zeolites thus treated are herein referred to as the stabilized hydrogen form of the zeolite. Sometimes only a portion of the remaining exchange capacity will be occupied by stabilized hydrogen ions. In those instances the remainder of the ion exchange capacity may be occupied by ions of another nature.
If desired, the resultant zeolite can be subjected to further ion exchange and steaming to increase the hydrogen ion content and correspondingly reduce the alkali metal content. However, I have found that the necessary degree of exchange can be efficiently accomplished by one exchange-steaming cycle.
The resultant steamed zeolite is then reexchanged with a hydrogen ion precursor under conditions sufficient to reduce the alkali metal content to less than 2 percent, usually less than one percent and preferably less than 0.6 weight-percent determined as the corresponding alkali metal oxide. These conditions are usually sufficient to produce a zeolite containing an amount of hydrogen precursor ion corresponding to at least about 5 relative percent of the original ion exchange capacity of the aluminosilicate.
Although this treatment can be applied to a variety of aluminosilicates, it is presently preferred to employ as starting materials a composition that contains at least a substantial proportion of a faujasite type of zeolite similar to the Y-zeolite described in U.S. Pat. No. 3,130,007. In the sodium form these zeolites usually contain pores in the range of about 5 to about 16 angstroms diameter and have relatively uniform pore size distributions. However, I have found that by subjecting those zeolites to the above-described preactivation that several beneficial changes in chemical and physical characteristics take place. Of these probably the most significant are increased activity, selectivity to products boiling within a predetermined range, thermal, hydrothermal, acid and reammoniation stability and increased pore size. With regard to this latter consideration I have discovered that this procedure accounts for a broadening of the pore size distribution with the result that, in the case of Y-zeolite type starting materials, at least about 20 percent of the pore volume of the zeolite is contained in pores having diameters in excess of about 20 angstroms. Although it has not been established with certainty, it is believed that this increase in pore size may account for some of the observed improvements in selectivity of the final catalyst compositions. Accordingly, it is presently preferred that the zeolites have non-uniform pore size distributions. In particular, it is believed that at least about 40% of the zeolites pore volume should be accounted for by pores having diameters of less than about 20 angstroms and that at least about 20 percent should be accounted for by pores having diameters in excess of about 20 angstroms.
Other zeolites preferred for use in these midbarrel hydrocracking processes are the so-called ultrastable and layered zeolites referred to above.
The refractory oxide must be precalcined at a temperature of at least about 600.degree. F., preferably about 800.degree. to about 1400.degree. F., prior to combination with the zeolite and hydrogenation components. Calcination generally requires at least about 20 minutes, preferably about 30 minutes to about 6 hours.
A number of refractory oxides can be employed in these compositions. These oxides should have relatively high surface areas, e.g., above about 50 square meters per gram, should be compatible with the zeolite and hydrogenation components and should combine with the zeolite to form structurally stable aggregates. These oxides include alumina, titania, zirconia and silica-magnesia, either alone or in combination with each other and/or other oxides, e.g., silica-alumina-magnesia, silica-titania, and the like. Alumina is presently preferred, particularly when high selectivity to a certain product fraction is desired. Although it can be used in combination with other oxides such as those mentioned, the alumina should account for at least 20, usually at least 50, and preferably at least 70 percent of the refractory oxide. Minor amounts of silica can be added to these preferred catalysts although it is preferably present in amounts less than about 20 weight percent. Silica-magnesia is also useful in preparing these highly selective compositions. The combination usually contains about 5 to about 40 percent magnesia based on the combined weight of silica and magnesia.
In addition, the refractory oxide, preferably alumina, silica-alumina, silica-magnesia, or combinations thereof should not be peptized by acid treatment to any substantial extent or otherwise hydrolyzed before admixture with said zeolite and said hydrogenation component. It is sometimes desirable to peptize a minor portion of the alumina in such compositions to improve aggregate strength. However, the extent of such treatment should be kept to a minimum so that less than 50 percent of the oxide is rehydrolyzed after the thermal pretreatment.
Alumina sources include a variety of dried or hydrous gels, sols, spray-dried aluminas and the like. However, the boehmite and gamma forms are presently preferred.
I have discovered that the described treatments result in the production of compositions in which at least about 70 percent of the alumina pore volume is in pores having diameters above about 50 angstroms. Conversely about 50 percent of the pore volume is accounted for by pores having diameters less than about 200 angstroms. It has not been established that these pore volume distributions account for any part of the noted improvements in activity and selectivity. However, it is presently believed that they account, at least in part, for one or more of the noted improvements in catalyst performance.
The third essential component of these compositions, the hydrogenation component, usually comprises one of the metals, oxides or sulfides of Groups VI and VIII. The presently preferred compositions include at least one of cobalt and nickel oxides or sulfides and at least one of molybdenum and tungsten oxides and sulfides. The advantages of these methods are particularly apparent with compositions containing nickel and/or cobalt sulfide and molybdenum sulfide. Moreover, the greatest relative advantage is obtained when one or more of the hydrogenation components, particularly tungsten and molybdenum, and especially molybdenum, are combined with the alumina and/or zeolite prior to any high temperature thermal treatment. Accordingly, the hydrogenation component or precursor is usually combined with the refractory oxide and zeolite by intimate admixture of the three components in the presence of sufficient water to produce a formable plastic mixture. This approach is preferred due to the desirability of forming the resultant combination into particulate aggregates such as extrudates, tablets, spheres or the like. About 30 to about 60 weight percent water is usually sufficient for this purpose.
The Group VIII components, notably nickel and cobalt, can be added as water-soluble compounds such as the carbonates, sulfates, nitrates or halides. In the alternative, they can be present in the form of complex molybdenum salts, such as the complex cobalt or nickel molybdophosphates, molybdosilicates and the like. Similarly, the Group VI components, particularly tungsten and/or molybdenum, can be added as either soluble or insoluble compounds including the oxides, e.g., tungstic oxide, molybdic oxide, molybdenum blue, or salts such as ammonium phosphomolybdate, ammonium molybdate, ammonium dimolybdate and the complex metal salts mentioned above.
A particularly preferred method of combining these several components involves comulling the refractory oxide and zeolite with at least one nickel or cobalt compound and at least one molybdenum or tungsten compound in the presence of water. The aqueous medium can also contain a constituent capable of solubilizing the Group VI and/or Group VIII components or complexing the Group VI and Group VIII components with each other. Exemplary materials are orthophosphoric acid, ammonium hydroxide, and hydrogen peroxide, orthophosphoric acid being presently preferred. Thus, the aqueous phase can contain sufficient orthophosphoric acid to reduce the pH to a level below about 5, preferably below about 4, and introduce at least about 0.5 and preferably about 1 to about 7 weight percent phosphorus as P.sub.2 O.sub.5. The catalyst usually contains at least about 0.5 and preferably about 2 to about 10 weight percent of the Group VIII component and at least about 1, preferably about 2 to about 40 weight percent of the Group VI component determined as the respective oxides.
The refractory oxide and zeolite should be admixed in finely divided particulate form such that an intimate dispersion of each component with the other can be easily achieved. Accordingly, it is preferred that the predominance of both the oxide and zeolite be in the form of particles, powders, flakes or the like having average diameters of less than about 2 microns.
The amount of zeolite thus added should correspond to at least about 1, and preferably about 2 to about 80 weight percent based on the total dry weight of the zeolite and oxide. However, the compositions presently most preferred for denitrogenation and/or selective hydrocracking usually contain about 2 to about 30 weight percent zeolite.
The intimate dispersion of refractory oxide, zeolite and hyrogenation components is then formed into the desired particle aggregate by any one of the several known procedures including extrusion, pelleting and the like. The pellets are then calcined directly with or without initial drying. Calcination temperatures are usually in excess of about 600.degree. F., preferably at about 800.degree. to about 1500.degree. F. The calcined composition is then preferably sulfided by contacting with a sulfur donor for a period sufficient to convert the hydrogenation metals or metal oxides to the corresponding sulfides. Conventional sulfur donors include hydrogen sulfide, carbon bisulfide, elemental sulfur, hydrocarbon thiols and thioethers having up to 10 carbon atoms per molecule, and the like. In the alternative, the catalyst can be sulfided in situ in a hydrocarbon conversion zone by exposure to a hydrocarbon feed containing organosulfur compounds under conditions sufficient to convert the metals or metal oxides to the corresponding sulfides.
In general these processes involve the reaction of hydrocarbons with elemental hydrogen under hydroconversion conditions of temperature, pressure and contact times sufficient to react at least about 50 standard cubic feet of hydrogen with each barrel of hydrocarbon feed. However, these compositions exhibit the greatest advantage in processes of more limited scope. These include the hydrogenative conversion of hydrocarbons to lower molecular weight products boiling within predetermined boiling ranges and hydrofining systems, particularly those involving hydrogenative denitrogenation.
In most hydrocracking processes a principle portion of the feed boils in excess of about 300.degree. F., usually in excess of about 500.degree. F. However, in accordance with a preferred embodiment midbarrel fuels boiling primarily between about 350.degree. and about 700.degree. F. are selectively produced from feeds boiling primarily above about 700.degree. F., usually about 700.degree. to about 1300.degree. F. Usually at least about 70 percent of the feed in the preferred midbarrel systems will boil above 700.degree. F. Exemplary refinery feedstocks are straight run gas oils, vacuum gas oils, deasphalted vacuum and atmospheric residua, coker distillates, catcracker distillates, cycle stocks and the like.
Hydrocracking systems are distinguished from other hydrogenative reactions such as aromatics and olefin hydrogenation, denitrogenation and desulfurization, by the substantial reduction in initial boiling point of the hydrocarbon feed. For the purposes of this invention, hydrocracking involves the conversion of at least 20 volume-percent of the feed to materials boiling below its initial boiling point. In most commercial applications it is generally preferred to convert at least 40 volume-percent of the feed per pass.
At times, however, hydrocracking cannot be characterized in this manner due to the inclusion of minor amounts of relatively low boiling materials in the feedstock. Nevertheless it may be distinguished from less severe hydrogenative processes by comparing the number of moles of product produced to the amount of feedstock reacted. On this basis hydrocracking usually involves the production of at least 110 moles of product for each 100 moles of feed. However, higher conversions involving the production of at least 120 moles of product for each 100 moles of feed are generally preferred. These reactions can be even further characterized by relatively higher net hydrogen consumption which usually exceeds about 250 standard cubic feet net hydrogen consumed per barrel of feed.
As illustrated by the examples hereinafter detailed, the compositions and methods of this invention are particularly attractive for the conversion of higher boiling hydrocarbons to either midbarrel or gasoline range products. The selectivity of these systems for the desired product of predetermined boiling range, i.e., midbarrel or gasoline products, is vastly superior to alternative systems.
Considerable overlap can and does exist between the definitions of midbarrel and gasoline range hydrocarbons. At least part of this overlap depends on the selection of product cut points for convenience of identification. A much more significant variable however is the difference in product properties required to meet specific end uses, and/or the tailoring of hydrocrackate required to obtain the optimum performance of post treatment systems such as reforming and isomerization. However, as a general rule midbarrel products, a category which includes diesel fuels, turbine fuels and furnace oils, are usually characterized by a boiling point range of about 300.degree. to about 700.degree. F. Diesel fuels boil primarily below about 570.degree. F. while turbine and furnace fractions boil predominantly below 675.degree. F. and 700.degree. F., respectively. The C.sub.5 to about 500.degree. F. fraction is generally classified as gasoline.
Regardless of the exact definition given to the product fraction, any attempt to produce any of these products involves conversion of a substantial proportion of the feed to hydrocarbons boiling below the desired product range. If the amount of such conversion can be reduced while maintaining the same rate of conversion to the desired product boiling range, the economics of the system are obviously improved. It is this balance of overall conversion and conversion to the desired product, that is referred to as selectivity. It is therefore significant that these methods exhibit a greater degree of selectivity to a specified product boiling range than do analogous processes. In other words, a greater proportion of the product boils within the desired predetermined range. Conversion to products boiling outside this range, particularly to lower boiling materials, is correspondingly reduced.
An additional advantage of these methods is that their activity and selectivity is affected to a much lesser extent by catalyst inhibitors such as organonitrogen compounds. Accordingly, the hydrocarbons employed in these systems often contain in excess of about 5, and can contain as much as 50 or 900 ppm or more of nitrogen as organonitrogen compounds without losing too much efficiency.
Conversion is usually carried out at temperatures of at least 400.degree. F. preferably in excess of 600.degree. F., and generally in the range of about 600.degree. to about 900.degree. F. Reaction pressures are generally over 500 psig. in hydrocracking systems although lower pressures can be employed for denitrogenation. However, most commercial hydrocracking operations involve pressures in excess of about 100 psig., usually about 1500 to about 3000 psig.
The degree of contacting of the hydrocarbon feed at these conditions will of course depend upon the extent of conversion desired and the severity of reaction conditions required to obtain that conversion. In essentially every case, including both hydrocracking and denitrogenation, contact times will exceed one minute. Corresponding liquid hourly space velocities are usually less than about 10 and are preferably within the range of about 0.3 to about 5. Hydrogen addition rates correspond to about 500, and, usually about 4000 to about 20,000 SCF of hydrogen per barrel of hydrocarbon.
As a general rule when hydrocracking to midbarrel fuels is a primary objective, reaction conditions within the ranges discussed will usually be selected so as to effect at least about 40 percent conversion per pass to products boiling below the initial feed boiling point, e.g., less than about 700.degree. F. Conversion efficiency is usually indicated by a selectivity of at least about 50 percent.
A somewhat wider range of reaction conditions and feedstock characteristics can be employed in denitrogenation processes. For example, there are essentially no limitations on the organonitrogen content of the feed. Nitrogen levels can range from 5 ppm up and are usually about 50 ppm to about 1.5 percent. The feedstock boiling range or initial boiling point is also much more variable than in the case of gasoline or midbarrel hydrocracking. For example, these materials can boil as low as 100.degree. F. but generally boil above about 400.degree. F. The ranges of reaction temperatures, pressures and space velocities are quite similar to those discussed with regard to hydrocracking. However, lower pressures and higher space velocities are generally sufficient to accomplish the desired degree of denitrogenation without substantial hydrocracking.